Swing-bed guard chamber in hydrogenerating and hydrorefining coke-forming hydrocarbon charge stock

ABSTRACT

A PROCESS FOR HYDROREFINING SULFUROUS HYDROCARBON DISTILLATES CONTAINING MONO-OLEFINIC, DI-OLEFINIC AND AREMATIC HYDROCARBONS. PARTIULARLY DIRECTED TOWARD THE STABILIZATION OF A PYROLYSIS NAPHTHA, THE PROCESS MAKES USE OF A MULTIPLE-STAGE SYSTEM. EASE OF OPERATION, INCREASED CATALYST STABILITY THROUGHOUT THE PROCESS AND VARIOUS ECONOMIC ADVANTAGES ARE OBTAINED THROUGH THE UTILIZATION OF A SWING-BED REACTION ZONE SYSTEM WHICH CONSTITUTES THE INITIAL PORTION OF THE PROCESS. IN COMBINATION WITH THE   FIRST PRINCIPAL REACTION ZONE, FOLLOWING THE SWING-BED SECTION, EACH OF THE SMALLER GUARD CHAMBER REACTORS IN THE LATTER CONTAIN FROM 5.0% TO ABOUT 25.0% OF THE TOTAL CATALYST.

Oct. 3, 1972 SWING-BED GUARD CHAMB'ER L E. HUTCHINGS IN HYDROGENATING`AND HYDROREFINING A COKE-FORMING HYDROCARBON CHARGE STOCK Filed July27, 1970 A TTORNEYS Unted States Patent O 3,696,022 SWING-BED GUARDCHAMBER IN HYDRO- GENERATING AND HYDROREFINING COKE- FORMING HYDROCARBONCHARGE STOCK Leroi E. Hutchings, Mount Prospect, Ill., assignor toUniversal Oil Products Company, Des Plaines, Ill. Filed July 27, 1970,Ser. No. 58,443 Int. Cl. Cg 23/00 U.S. Cl. 208--57 9 Claims ABSTRACT OFTHE DISCLOSURE APPLICABILITY OF INVENTION In the present specificationand appended claims, the terms hydrocarbons, hydrocarbon fractions,hydrocarbon distillates, and hydrocarbon mixtures are utilizedinterchangeably to connote synonmously mixtures of hydrocarbonsresulting from diverse conversion processes. Such processes include thecatalytic and/ or thermal cracking or petroleum, the latter oftenreferred to as pyrolysis, the destructive distillation of wood or coal,shale oil retorting, etc. The resulting hydrocarbon distillate fractionscontain impurities which must necessarily be removed before thefractions are suitable for their intended use, or, when removed, enhancethe value of the fraction with respect to further processing.Contaminating influences, in the form of these impurities, includesulfurous compounds, nitrogenous compounds and often oxygenatedcompounds.

In addition to the aforementioned contaminating inuences, thesehydrocarbon fractions contain appreciable quantities of unsaturatedcompounds, including monoolefinic, di-olefinic (including `conjugateddi-olefns) and aromatic hydrocarbons. It is the mono-olefinic anddiolenic hydrocarbons, particularly butadiene and pentadienes, whichinduce the polymer and coke-forming tendencies of the hydrocarbondistillate, as do the vinyl aromatics, and which, when the distillate issubjected to hydrotreating for the purpose of removing the contaminatinginfluences, effect the formation of polymers as Well as othercarbonaceous material. The formation of olefin polymers, and depositionof coke, appears to be a result of the necessity for effecting thehydrotreating process at f elevated temperatures above about 500 F. inorder to convert sulfurous compounds into hydrogen sulfide andhydrocarbons. Not only are the hydrotreating catalytic compositesaffected, but various heaters and other appurtenances of the conversionzone experience heavy coking which appears as a formation of solid,highly carbonaceous material principally resulting from thepolymerization of conjugated di-olefns, and styrenes and olens containedin the fresh feed charge to the unit.

As hereinbefore set forth, coke-forming hydrocarbon distillates areusually those fractions resulting from prior conversion processes suchas catalytic or thermal cracking, or destructive distillation. In theinterest of brevity, the following discussion will be directed towardhydro- 3,696,022 Patented Oct. 3, 1972 ice carbon distillates resultingfrom naphtha pyrolysis units generally designed for the production ofnormally gaseous olefinic material such as ethylene, propylene,butadiene, etc. Pyrolysis naphtha co-product is available in relativelylarge quantities, but requires a hydrorefining treatment for the purposeof enhancing its possibilities for further usefulness. In manyinstances, the pyrolysis naphtha c0- product Will not contain excessivequantities of sulfurous and/or nitrogenous compounds, but will consistof detrimental amounts of mono-olefins and di-olens, includingconjugated di-olens, to the extent that immediate use of the distillateis prohibited. For example, in a thermal cracking process designed forethylene production, a full boiling range hydrocarbon distillate isproduced which may contain less than 1,000 p.p.m. by weight each ofsulfur and nitrogen, but sufiicient quantities of olefinic hydrocarbonsto indicate a bromine number of the order of at least about 25.0, andoften more, and di-olens in an amount to indicate a diene value of theorder of about 20.0, or more.

Pyrolysis reactions are effected in the absence of a catalyticcomposite, at elevated temperatures and in the presence of a diluentsuch as superheated stream. Depending upon the physical and/or chemicalcharacteristics of the charge stock, as well as the specific pyrolysisconditions, the product eiuent comprises varying quantities of lightoleinic hydrocarbons including ethylene, propylene, `butylene,butadiene, etc., and a pyrolysis naphtha fraction containing pentanes,hexanes and heavier hydrocarbons boiling up to a temperature in therange of about 300 F. to about 400 F., and including aromatichydrocarbons, mono-olefnic and di-olefinic hydrocarbons, styrenes andsulfurous compounds. Visualized as being exemplary of those pyrolysisnaphthas intended for processing, in accordance With the presentinvention, are: a naphtha having a gravity of about 35.1 API, andcontaining 1,100 p.p.m. by weight of sulfur, having an end boiling pointof 350 F., a bromine number of 43.0 and a diene value of 40.0; a 400 F.(end boiling point) naphtha having a gravity of 43.5 API, and containing500 p.p.m. by Weight of sulfur, having a bromine number of 74.0 and adiene value of about 80.0; and, a butane/ pentane concentrate having agravity of 76.7 API containing 500 p.p.m. by Weight of sulfur, andhaving a lbromine number of about 200.0 and a diene value of about230.0.

Since such pyrolysis naphtha fractions are severely contaminated, it isa common practice to hydrotreat or saturate the olens and/or di-olens,while destructively converting the sulfurous and nitrogenous compoundsinto hydrogen sulfide, ammonia and hydrocarbons. As hereinbefore stated,di-olefnic hydrocarbons present particular diiculty in the operation ofthe hydrotreating facilities by way of catalyst deactivation andextensive equipment fouling. Attempts are often made to improve theon-stream efliciency of the hydrotreating process by either promotingthe polymerization of the diolenic hydrocarbons prior to thehydrotreating step, or by utilizing operating techniques which tend tominimize or inhibit polymer formation. However, none of these approachesare suiiiciently successful in overcoming the difficulty, and to theextent that the process enjoys an acceptable degree of catalyststability.

Of further significance is the lack of selectivity in prior artprocesses. For example, the hydrogenation of the conjugated di-oleiinichydrocarbons and styrenes may not cease with the conversion thereof tosaturates and alkyl benzenes, but will continue to complete saturationof the mono-olens in the charge stock and the aromatic nuclei in thestyrene compounds. Such non-selectivity obviously results in a decreaseof desirable products in the thus-treated product efuent. Candor compelsrecognition of the fact that usch prior art processes have recently beenimproved. Since pyrolysis naphtha fractions tend to form gums and otherolefin polymer material upon standing, one such improvement involvesremoval of this material, butadienes and pentadienes from the fresh feedcharge stock in a fractionation column; this technique is commonlyreferred to in the art as rerunning and depentanizing the fresh feed.Additionally, another principal improvement involves the use of amultiple-stage system in which the first stage functions at atemperature below that at which desulfurization is effected, or belowabout 500 F., while the second stage, in order to effectdesulfurization, functions at temperatures exceeding 500 F.Nothwithstanding these improvements, which admittedly result inincreased catalyst stability and improved product quality additionalimprovements are afforded through the utilization of my invention. Thepresent process improves the quality of the matreial being charged toboth of the principal reaction zones, thereby further increasingcatalyst stability, or catalyst life. Ultimate product quality is alsofurther improved with respect to the quantity of mono-olefinichydrocarbons. Heretofore, particularly in European and other foreignlocales, the presence of mon-olefinic hydrocarbons, as a result of theirexcellent blending values, were tolerated in motor fuel gasolines. Sincethe burning of olefin-containing motor fuel has been found to be acontributing factor to pollution of the atmosphere, stringentrequirements are imposed whereby the olefin concentration of motor fuelgasoline must be essentially nil-77 In accordance with the presentprocess, through the use of particular catalytic composites, operatingconditions and especially operating techniques, a coke-forminghydrocarbon distillate is selectively hydrogenated and desulfurized withminimum polymer formation and minimum degradation of desiredconstituents. Briefly, the present invention provides a series of guardchambers, in the form of catalytic reaction zones, employed in aswing-bed arrangement whereby one or more chambers are on-stream, whileone or more chambers are simultaneously being regenerated. This featureof the present invention eliminates the necessity for re-running of thefresh feed charge stock and improves catalyst stability, and is basedupon recognition of the fact that polymer formation, notwithstandingoperating temperatures below about 500 F., takes place in the initial5.0% to about 25.0% of the first stage catalyst bed. This feature leadsto a continuous hydro-treating process through the elimination offrequent shut-down periods for regeneration of the catalytic composites.

In the absence of the guard chambers, but including the re-running anddepentanizing of the fresh feed charge stock, a commercial unit must beshut down for hot hydrogen stripping, to remove olefin polymers, everyten to sixty days, the latter at the start of the run, the former nearthe end of the run when the catalytic composite must be burned Throughthe incorporation f the guard chambers, and elimination of re-runningand depentanizing, shut downs for hydrogen stripping of the catalyst inthe principal reaction zones are nonexistent a requirement for burningthe catalyst in the guard chambers does not lead to a shut down, and thenecessity for burning the principal catalytic beds is greatly reduced.

OBJ-ECTS AND EMBODIMENTS A principal object of the present invention isto effect the saturation of olefinic hydrocarbons and thedesulfurization of sulfurous compounds present in coke-forminghydrocarbon distillates. A corollary objective is to provide catalyststability in a process for stabilizing cokeforming pyrolysis gasoline.

A specific object of my invention is to provide a continuoushydrotreating process into which catalyst regeneration facilities areintegrated with the result that catalyst stability is significantlyimproved and little degradation of aromatic hydrocarbons exists.

Therefore, in one embodiment, my invention provides a continuous processfor hydrorefining a sulfurous, hydrocarbon charge stock containingmono-olefinic and diolefinic hydrocarbons, which process comprises thesteps of: (a) reacting said charge stock and a regeneratinghydrogen-rich gaseous phase in a first catalytic reaction zone at amaximum catalyst bed temperature in the range of about 200 F. to about500 F.; (b) further reacting the resulting first reaction zone efiiuent,at substantially the same temperature in the range of about 200 F. toabout 500 F., in a second catalytic reaction zone; (c) increasing thetemperature of the resulting second reaction zone efiiuent to a levelabove about 500 F., and reacting the thus-heated second zone effluent ina third catalytic reaction at a maximum catalyst bed temperature belowabout l,000 F.; (d) separating the resulting third reaction zoneefiiuent, at a temperature in the range of about 200 F. to about 500 F.,in a first separation zone, to provide a first vaporous phase and afirst liquid phase; (e) further separating said first vaporous phase, ata substantially lower temperature in a second separation zone, toprovide a second liquid phase and a hydrogen-rich second vaporous phase;(f) heating said hydrogen-rich second vaporous phase to a temperature inthe range of about 500 F. to about 700 F.; (g) passing the thus-heatedsecond vaporous phase through a fourth catalytic reaction zonecontaining an olefin polymer-deactivated catalytic composite; (h)separating resulting fourth reaction zone efiiuent, in a thirdseparation zone, to provide a polymer concentrate and a hydrogen-richthird vaporous phase; and, (i) recyclying said third vaporous phase tocombine with said charge stock.

Other embodiments of my invention involve preferred processingtechniques and operating conditions, as well as particular catalyticcomposites for utilization in the various catalytic reaction zones. Withrespect to the latter, it is preferred to utilize, in the initial majorcatalytic reaction zone, a catalytic composite capable of selectivelyhydrogenating conjugated di-olefinic hydrocarbons in a hydrocarboncharge stock, also containing sulfurous compounds, aromatic hydrocarbonsand mono-olefinic hydrocarbons, to the corresponding saturates withoutsignificant saturation of the mono-olefins present. Thus, a preferredcatalyst for use in the first reaction zone is a composite of a GroupVIII noble metal component, an aluminacontaining non-acidic carriermaterial and an alkali metal component. With respect to the catalyticcomposite disposed within the second major reaction zone, intended fordesulfurization and mono-olefin hydrogenation, accompanied bypreservation of the aromatic hydrocarbons, the character of thecatalytic composite, as hereinafter indicated, is dependent upon themaximum catalyst bed temperature within the reaction zone.

In another embodiment, the first and fourth catalytic reaction zonesfunction in swing-bed fashion and are interchanged when the deactivatedcatalytic composite in the fourth reaction zone is substantially freefrom olefin polymers, with the result that the charge stock andregenerating hydrogen gaseous phase are then introduced into the fourthreaction zone, and the heated hydrogen-rich second vaporous phase isintroduced into the first catalytic reaction zone to remove olefinpolymers from the catalyst disposed therein.

SUMMARY OF THE INVENTION As hereinbefore set forth, the principalpurpose of the present invention is to provide a selective and highlystable continuous process for hydrogenating coke-forming hydrocarbondistillates. As utilized herein, the term hydrogenating is intended tobe synonymous with hydrotreating and hydrorefining. In essence, thispurpose is accomplished through the use of a two-stage fixed-bedcatalytic reaction system, having integrated therein a swingbedregeneration technique.

Previously, there existed two separate, desirable routes for thetreatment of coke-forming distillates. One such route was directedtoward recovering a hydrocarbon product suitable for use in certaingasoline blending. With this as the desired object, the hydrorefiningprocess could be effected in a. single-stage, or reaction zone, ingeneral utilizing the catalytic composite hereinafter specificallydescribed as the third reaction zone catalyst. The attainableselectivity in this instance resides primarily in the hydrogenation ofhighly reactive double bonds, while leaving native mono-olefins in thecharge stock unreacted; with respect to styrene, the hydrogenation isinhibited to produce alkyl benzenes without ring saturation. Until thestringent restrictions were placed upon the olefin content of motorfuels, the single stage process was suitable for use in some locales.Where, however, the desired end result was aromatic hydrocarbonretention, intended for subsequent recovery via extraction, thetwo-stage route was required. Currently, regardless of the ultimate useof the aromatic-containing product, the two-stage route becomesnecessary. It has previously been noted that the sulfurous compounds andthe mono-olefins remain substantially unchanged in the single, or thefirst-stage reaction zone. This restrictive hydrogenation selectivity isaccomplished with reduced polymer formation to gums Other advantages ofrestricting the hydrogenation to the conjugated di-olefins and theolefinic bond in the styrenes include: lower hydrogen consumption, alower heat of reaction and, obviously, increased aromatic retentionincluding the alkyl benzenes resulting from styrene hydrogenation. Themono-olefins must be saturated in the second stage, not only to`facilitate aromatic extraction by way of currently known and utilizedmethods, but also to comply with the impositions placed upon the olefincontent of motor fuels. Thus, the desired, necessary second-stagehydrogenation technique involves saturation of the monoolefins, as wellas sulfur removal, the latter required for an acceptable aromaticproduct.

As hereinafter set forth, in the description of the accompanyingdrawings, the present continuous process requires more than twocatalytic reaction zones, and at least four catalytic reaction zones.Two initial reaction zones serve as swing-bed guard chambers, each ofwhich, when considered in conjunction with the first major, or principalreaction zone, contains about 5.0% to about 25.0% by weight of the totalcatalytic composite disposed in both zones. In the present specificationand appended claims, the first and fourth catalytic reaction zones arethe guard chambers, and generally contain the same catalytic compositeas is preferably disposed within the second catalytic reaction zone. Byway of clarification, the second and third reaction zones of the presentinvention are those to which the prior art refers to as the first andsecond stages, respectively. In most applications of the presentprocess, the catalytic composite will be the same in the first, secondand fourth catalytic reaction zones.

The principal function of the catalytic composite utilized in the secondreaction zone, and preferably in the first and fourth reaction zoneguard chambers, involves selective hydrogenation of conjugateddi-olefinic hydrocarbons. One particular catalytic composite possessesunusual stability notwithstanding the presence of sulfurous compounds inthe fresh hydrocarbon charge stock. The catalytic composite ishereinafter described in greater detail. Briefly, however, the catalystis a composite of an alumina-containing, non-acidic carrier material, aGroup VIII noble metal component and an alkali metal component, thelatter being preferably lithium and/ or potassium. In order to avoidhydrocracking activity, it is particularly preferred that this catalyticcomposite be substantially free from any acid-acting propensities. Withrespect to the third catalytic reaction zone, the catalytic compositefunctions for the principal purpose of effecting Ithe destructive Citconversion of sulfurous and nitrogeneous compounds, as well asmono-olefin saturation, and is a composite of an alumina-containingcarrier material, a Group VIII metal component; it may or may not alsocontain a halogen component. As hereinafter set forth, the character ofthe Group VIII metal utilized in the third catalytic reaction zone, isconsidered in conjunction with the maximum catalyst bed temperatureemployed therein.

Through utilization of a particular sequence of processing steps, theessence of which involves an integrated regeneration technique when hothydrogen, employed as the regenerating gas, is introduced into theprocessin-g section, catalyst stability is improved to a degree whichaffords a continuous process and makes it unnecessary to rerun ordepentanize the feed. Regeneration of the catalyst disposed in the majorprocessing reactors can thus be postponed until such time as the unit isscheduled for its normal maintenance shut down. This sequence ofprocessing steps, hereinafter set forth in greater detail, regulates theoverall hydrorefining process in such a manner that the charge to themajor processing section is not at conditions which induce the olefinpolymerization reactions. The present process provides an improvement inthe quality of charge stock to both the second and third catalyticreaction zones.

PROCESS CONDITIONS AND OPERATIONS The process of the present inventionis effected in a sequence of reaction zones, each of which is maintainedat operating conditions consistent with the chemical characteristics ofthe reactants passing therethrough. In order to afford a clearunderstanding of the integrated process, brief reference to theaccompanying drawing is believed warranted. In the drawing, reactors 12and 18 are those which have previously been referred to as guardchambers. For the purpose of the following discussion, reactor 12 willbe considered as undergoing regeneration with a hot hydrogen-richgaseous phase, whereas reactor 18 is operating within the process.Throughout the specification, reactor 18 is referred to as the firstreaction zone and reactor 12, the fourth reaction zone. The principalcatalytic reaction zones are reactors I7 and 27, and are herein referredto as the second and third reaction zones, respectively. With respect tothe distribution of the catalytic composite in the first, second andfourth reaction zones, each of the first and fourth zones will containfrom about 5.0% to about 25.0% of the total catalyst disposed in eachzone in conjunction with the quantity of catalyst disposed in the secondcatalytic reaction zone. That is to say, the distribution of catalystwith respect only to the first and second reaction zones is such thatthe quantity of catalyst in the first reaction zone is Within theaforesaid range.

Processing some charge stocks, having an extreme degree of unsaturationand high concentrations of contaminating influences, may result in toogreat a rise in temperature due to the exothermicity of the reactions.In such instances, a desirable technique constitutes providing formulti-point introduction of either the liquid feed, or internallyrecycled diluents, or recycled hydrogen-rich gaseous phase at variousintermediate loci of the reaction zones. This tends to prevent excessivesaturation from occurring in one particular portion of the catalyst, andalso cools the charge stream as it passes through the reaction zone.

A hydrogen-rich recycle gaseous phase, the source of which ishereinafter set forth, is heated to a temperature in the range of about500 F. to about 700 This hot hydrogen-rich stream is utilized to stripthe catalytic composite which has become deactivated as a result of theformation of olefin polymers. The regenerating effluent is introducedinto a suitable separation zone, such as knock-out pot, from which thestripped polymers and gums are withdrawn. Following its use as aheat-exchange medium, to reduce its temperature, the

hot regenerating hydrogen stream is combined with the hydrocarbondistillate charge stock, and liquid diluent where used, the mixturebeing introduced into the rst operating catalytic reaction guardchamber. The hydrocarbon distillate charge stock, for example a lightnaphtha by-product from a commercial thermal cracking unit designed andoperated for the production of ethylene, having a gravity of about 39.9API, a bromine number of about 42.0, a diene value of about 32.0 andcontaining about 100 p.p.m. of sulfur, 100 p.p.m. of nitrogen, and about71.5 vol. percent aromatic hydrocarbons, is admixed with an internallyrecycled liquid diluent, the source of which is hereinafter set forth,to provide a combined liquid feed ratio in the range of about 1.1 toabout 6.0. The hydrogen concentration is within the range of from about500 to about 10,000 scf/bbl., and preferably in a narrower range ofabout 1,000 to about 6,000 sci/bbl. The maximum catalyst bed temperatureis maintained in the range of about 200 F. to about 500 F., andpreferably at a level of about 250 F. This guard chamber functions undera pressure of from about 100 p.s.i.g. to about 1,000 p.s.i.g., andpreferably at a level in the range of from about 500 p.s.i.g. to about900 p.s.i.g.

Without substantial change in either temperature, or pressure, theeliiluent from the guard chamber is introduced into a second catalyticreaction zone, being the rst of the two major reaction zones. Thissecond reaction zone preferably contains a catalyst of the samecharacter as that in the guard chamber, being a composite ofalumina-containing non-acidic carrier material, a Group VIII noble metalcomponent and an alkali metal component. As hereinbefore set forth,considering only the operating guard chamber and this second catalyticreaction zone, the former contains from about 5.0% to about 25.0% of thecatalyst disposed in both zones.

The product efuent from the second catalytic reaction zone is introducedinto a suitable direct-tired heater wherein the temperature is increasedto a level above about 500 F. and preferably in the range of about 550F. to about l000 F. When the process is functioning eiciently, the dienevalue of the normally liquid hydrocarbonaceous material entering thethird catalytic reaction zone is in the range of about 0.1 to about 0.5,and very often nil. The conversion of nitrogenous and sulfurouscompounds, and the saturation of mono-oletins, contained within thesecond reaction zone eluent, is effected in the third catalytic reactionzone. The second catalytic reaction zone is maintained under an imposedpressure of from about 100 to about 1,000 psig., and preferably at alevel of from about 500 to about 900 p.s.i.g. Process operations arefacilitated when the focal point for pressure control is the coldseparator following the third catalytic reaction zone and, therefore',it will be maintained at a pressure slightly less than that imposed uponthe rst and second catalytic reaction zones. The third catalyticreaction zone and, therefore, it will be maincatalyst sufficient toprovide an LHSV of from about 0.5 to about 10.0, based upon fresh feed.The hydrogen concentration will be in the range of from about 500 toabout 10,000 s.c.f./bbl., based upon fresh feed, and preferably fromabout 1,000 to about 8,000 s.c.f./bbl.

Following its use as a heat-exchange medium, to decrease its temperatureto a level in the range of about 200 F. to about 500 F., and preferablyfrom about 250 F. to about 400 F., the product eluent from the thirdcatalytic reaction zone is introduced into a hot separator. A hot liquidphase is withdrawn from the hot separator and recycled to combine withthe fresh feed charge stock to the tirst catalytc reaction zone ashereinbefore described. Where desired or necessary, a portion of the hotseparator liquid may be employed to quench the reaction in the thirdcatalytic reaction zone by being introduced thereto at some intermediatelocus. The remainder of the hot separator liquid is then introduced intoa suitable separation system which removes hydrogen sulfide and recoversthe aromatic-rich naphtha boiling range fraction. The hydrogen-richvaporous phase from the hot separator is cooled and condensed to atemperature of from 60 F. to about 140 F., and passed into a suitablecold high pressure separator. One suitable operating technique involvesinjecting water into the hot vaporous phase from the hot separator andequipping the cold separator with a water dip-leg from which sour Watercontaining ammonia and hydrogen sulfide is removed. A hydrogen-richrecycle gaseous phase is withdrawn from the cold separator, and, by wayof compressive means, is recycled through a heater and a bed of olefinpolymer deactivated catalyst as hereinabove set forth. Make-up hydrogen,to supplant that consumed in the overall process may be introduced fromany suitable external source into any suitable location within theprocessing system; preferably, the make-up hydrogen is introduced by wayof the efuent line from the second catalytic zone.

The gaseous phase from the cold separator may be treated in any mannerfor the removal of hydrogen sulfide and/ or light paraflinichydrocarbons, in order to increase the concentration of hydrogentherein. The principally liquid phase from the hot separator may beadmixed with the liquid phase from the cold separator the mixture beingfurther separated in a hydrogen sulfide-stripping zone, the bottomfraction from which, as previously stated, represents the normallyliquid product of the process. With respect to the naphtha boiling rangeportion of the product effluent, the sulfur concentration is about 0.1p.p.m. by weight, the aromatic concentration is about 71.1% by volume,the bromine number is about 0.1 and the diene value is essentially nil.

In some situations, it may be desirable to recycle a portion of the hotseparator liquid to the inlet of the heater employed to increase thetemperature of the hydrogen-rich gaseous phase required to strippolymerization material from the catalyst in the guard chamber. Thisrecycled diluent also facilitates the removal of polymer from thecatalyst. In another embodiment, a portion may be recycled directly tothe guard chamber undergoing stripping in order to cool the same priorto swinging it into the system. For the most part, however, the recycleddiluent is employed for the dual purpose of inhibiting the temperatureincrease within the reaction system and to provide a cleaner combinedfeed stock to the second and third reaction zones. Since the producteluent emanating from the third catalytic reaction is (l) substantiallycompletely saturated with respect to olenc hydrocarbons, (2) virtuallycompletely desulfurized, it constitutes the source of the diluentemployed in the present continuous hydrorening process. The techniquesof separating the effluent in the hot separator, at a substantiallyreduced temperature in the range of about 200 F. to about 500 F., andpreferably from about 250 F. to about 400 F., provides a hot recyclestream consisting essentialy of the higher-boiling liquid components ofthe third reaction zone etllucnt. Thus, the recycle diluent comprisesprincipally Cc-Cg aromatic hydrocarbons which are substantially freefrom olelins, dissolved hydrogen sulfide, ammonia and lighter parafnichydrocarbons. The benelits accruing as a result of utilizing the hotrecycle diluent, in addition to considerations involved with heaterduties and heat-exchangers, will be evident to those having expertise inthe art of petroleum refining technology. However, it must be noted thatthere is also effected a decrease in the quantity of mon-olens in thetotal liquid feed to the third catalytic reaction zone. Thus,excessively great temperature increases are avoided leading to increasedcatalyst stability and a more economical operation from the standpointof the size and character of quench facilities which might otherwise beprovided.

DESCRIPTION OF CATALYTIC COMPOSITES The catalytic composites employed inthe present process comprise metallic components selected from themetals,

and compounds thereof of Group VI-B, and I-A and VIII of the PeriodicTable. Thus, in accordance with the Periodic Table of the Elements, E.H. Sargent & Company, 1964, suitable metallic components include lithiumsodium, potassium, rubidium, cesium, molybdenum, tungsten, cobalt,nickel, ruthenium, rhodium, palladium, osmium iridium and platinum.

While neither the precise composition, nor the method of manufacturingin the catalyst is essential to my invention, certain considerations arepreferred. For example, since the charge stocks to the present processare generally naphtha boiling range fractions, and the desired normallyliquid product eiiluent is a naphtha boiling range fraction, it ispreferred that neither catalytic composite exhibit an excessive degreeof hydrocracking activity, under the operating conditions utilizedherein, to the extent that the naphtha boiling range material isconverted into lower-boiling, normally gaseous hydrocarbon products.Furthermore, since a principal object of the present invention residesin the retention of aromatic hydrocarbons, excessive hydrogenationactivity is to be avoided. Although an acidic function may beincorporated within the catalytic composite utilized in the thirdcatalytic reaction zone, in order to facilitate the destructiveconversion of nitrogeneous and sulfurous compounds, the catalyticcomposite disposed within the first, second and fourth reaction zonesis, for the most part, non-acidic. Thus, with respect to the latter, thecatalytically active metallic components are preferably combined with anon-siliceous substantially halogen-free carrier material such asalumina. Substantially halogen-free composite is one wherein a halogencomponent is not intentionally added and, where a halogen compound(chloropalladic acid) is utilized in the catalyst manufacturingprocedure, steps are intentionally taken to remove halogen from theresulting composite.

The catalytic composites disposed within the first, second and fourthcatalytic reaction zones are preferably of the same composition andcharacter. This catalyst might be said to serve a dual-function; thatis, it must be nonsensitive to the presence of sulfurous compounds atthe operating conditions employed, while at the same time capable ofeffecting the hydrogenation of conjugated dioleinic hydrocarbons to thecorresponding saturates, while simultaneously possessing a degree ofselectivity such that the mono-oleiins and aromatic hydrocarbons are notsubstantially saturated. A catalyst comprising an alumina-containinginorganic oxide, combined with a Group VIII noble metal component and analkali metal component, is very etlicient in carrying out the desiredoperations. In contrast to the catalytic compositie utilized in thethird catalytic reaction zone, hereinafter described, it is preferredthat this catalyst be substantially free from acid-acting components,especially a halogen component As hereinbefore stated, a halogencompound is often employed during one or more steps of the overallcatalyst manufacturing technique. For example, alumina is commonlyprepared by a method which involves digesting substantially purealuminum metal in hydrochloric acid, and the Group VIII noble metal isoften impregnated throughout the finished alumina through the use of,for example, chloropalladic or chloroplatinic acid. It is Well 'known tobe extremely diflicult to remove combined halogen from the iinishedcatalyst to a level lower than about 0.1% by weight. The presence ofthis halogen, which imparts undesired acidity to the catalyticcomposite, is countered and inhibited through the use of the alkalimetal component.

The carrier material may be prepared in any suitable manner, and may benaturally-occurring. Following its preparation, the carrier material maybe formed into any desired shape including spheres, pills, cakes,extrudates, powders, granules, etc. Neither the form, nor the method ofmanufacturing the carrier materials is considered to be an essentialfeature of my invention. One component of the catalyst disposed in thefirst, second and fourth reaction zones, is an alkali metal componentemployed for the purpose of attenuating the inherent acidity possessedby residual halogen, or by the carrier material itself. ,Suitable alkalimetals are selected from the group of lithium, sodium, potassium,rubidium, cesium and mixtures thereof, particularly preferred metalsbeing lithium and/or potassium, Regardless of the particular state inwhich the alkali metal component exists, the quantities thereof, fromabout 0.05% to about 1.5% by weight, are calculated as if this componentexisted in the elemental state. Generally, it is preferred toincorporate the alkali metal cornponent during the separation of acarrier material; therefore, the carrier material is often refered toas, for example, lithiated alumina.

At relatively low temperatures, the Group VIII noble metals possess thepropensity for effecting the virtually complete hydrogenation of thereactive di-oleiins and styrenes, the latter selectively to alkylbenzenes. The noble metal components, selected from the group ofruthenium, osmium, rhodium, iridium, palladium, platinum and mixturethereof, are utilized in an amount of from about 0.01% to about 2.0% byweight, calculated as if existing in the elemental state. Group VIIInoble metal components may be incorporated within the catalyticcomposite in any suitable manner including co-precipitation with thecarrier, ion-exchange or impregnation of the carrier material with asuitable water-soluble compound of the metal. Following theincorporation of the metallic components, for example, by way ofimpregnation, the carrier material is dried at a temperature of about200 F. to about 600 F., and subsequently calcined, in an atmosphere ofair at an elevated temperature of about 700 F. to about 1200 F.

With respect to the catalytic composite disposed within the thirdreaction zone, for the primary purpose of destructively removingsulfurous compounds through the conversion thereof to hydrogen suliideand hydrocarbons, it should be noted that this reaction zone functionsat a higher temperature level in the range of about 500 F. to about1,000 F.

When the catalytically active metallic component is a Group VIII noblemetal component, the maximum catalyst bed temperature will be in therange of about 800 F. to about 1,000 F. When an ion-group metalcomponent is employed, either lone, or in combination with a Group VLBmetal component, the lower temperature range of about 500 F. to about800 F. will be employed. Although the third reaction zone catalyticcomposite may be similar to that utilized in the lirst major conversionzone, it is also distinctly different therefrom. For example, an alkalimetal lcomponent is `generally not combined therewith, and a halogencomponent is often combined therewith. The halogen may be eitherfluorine, chlorine, iodine, bromine, or mixtures thereof, with fluorineand chlorine being preferred. When utilized, the halogen component willbe composited in such a manner as results in a iinal compositecontaining about 0.1% to about 1.5% by weight, and preferably from about0.4% to about 0.9% by weight, calculated on an elemental basis. Withrespect to the alumina-containing carrier material, the alumina may beadvantageously employed in and of itself, or in combination with minorquantities of silica or other refractory inorganic oxides. When combinedwith, for example, silica, it is preferred that the alumina/silicaweight ratio be within the range of from about 63/ 37 to about /10. Whenthe third reaction zone catalytic composite contains a Group VIII noblemetal component, it is also selected from the group of ruthenium,rhodium, palladium, osmium, iridium and platinum. Of these, a palladiumand/or platinum metallic component is especially preferred. Thiscomponent may exist within the final catalyst composite as a compound,including the oxide, sulfide, halide, etc., or in an elemental state.The Group VIII noble metal component generally comprises about 0.01% toabout 1.0% by weight of the final catalytic composite, calculated on thebasis of the element.

When an ion-group metal component, particularly nickel and/or cobalt, isutilized, it will generally be present in an amount of about 1.0% toabout 10.0% by weight, again calculated as the elemental metal. Whenutilized, the Group VI-B metal component, particularly molybdenum and/ortungsten will be incorporated in an amount of from about 4.0% to about20.0% by weight.

With respect to interchanging or swinging the two reaction zones servingas guard chambers, referred to as the first and fourth reaction zones(in the drawing, reactors 18 and 12, respectively), a preferredtechnique is to effect the interchange at such times as the catalystundergoing regeneration becomes substantially free from olefinpolymerization products. This point will be noted by monitoring thematerial entering the knock-out pot for separation and removal of thepolymer products from the hot hydrogen gaseous phase. When this streamindicates that the olefin polymer products have been stripped from thecatalyst, the guard chamber may `be interchanged with the first reactionzone within the process. Simultaneously, the hot hydrogen becomesdiverted through the guard chamber just removed from the integratedprocess system. When this catalyst is substantially free from olefinpolymer products, the interchange is again effected. Another schemeinvolves monitoring the di-olefin content of the first reaction zoneefiluent. As the di-olefin content increases, there is an indicationthat the catalytic composite is losing activity as the result of theformation of olen polymers.

DESCRIPTION OF DRAWINGS My invention, as directed toward themultiple-stage hydroening of coke-forming hydrocarbon distillates, maybe more clearly understood upon reference to the accompanying drawingwhich illustrates one embodiment thereof. In the drawing, various owvalves, control valves, instrumentation and start-up lines, coolers,pumps and/or compressors, heat-exchangers, etc., have either beeneliminated, or reduced in number; only those connecting lines necessaryfor a complete understanding of the continuous process are shown. Theuse of other miscellaneous appurtenances is well within the purview ofone having skill in the art of petroleum processing techniques.

The drawing will be described in conjunction with a commercially-scaledunit designed to effect the multiplestage catalytic hydrotreating of athermally-cracked pyrolysis gasoline. The charge stock has a gravity ofabout 45.0 API, an initial boiling point of about 132 F. and an endboiling point of about 350 F. The charge stock contains about 700 p.p.m.by weight of sulfur and about 50.0% by volume of aromatic hydrocarbons,and indicates a bromine number of about 60.0 and a diene value of about50.0. The fresh feed, in amount of about 2,614 bbL/day, enters theprocess by Way of line 1, and is admixed with a liquid recycle diluentfrom line 2, the latter in an amount of 3,721 bbl/day. The mixturecontinues through line 2 and, via line 3, containing valve 5, and line36, into reactor 18. In the present illustration, it will be presumedthat reactor 18 is in operating status within the system, whereas thecatalyst in reactor 12 is undergoing regeneration. Therefore, withrespect to the various valves in the manifolding attendant these tworeaction zones, the open valves are designated as 5, 9, 22, 14, and 7,with the closed valves being 4, 3, 20, 11 and 17.

A hydrogen-rich recycle gaseous phase in line 33 is introduced intoheater 35, wherein the temperature is increased to a level of about 690F. The heated hydrogen passes through line 36 and open valve 7 in line 6into reactor 12. Hydrogen and stripped olefin polymers from the catalystdisposed in reactor 12 are withdrawn by way of line 13, containing openvalve 14, and introduced thereby into knock-out pot 15. Polymer productsare withdrawn from the system by way of line 38 containing open valve39. The hot hydrogen is withdrawn by way of line 37 and, after coolingto a temperature of about 265 F., passes through line 10 and open valve9 into line 36 wherein it is admixed with the fresh hydrocarbon chargestock and liquid diluent. Guard chamber 18 functions at a maximumcatalyst bed temperature of about 270 F., a liquid hourly spacedvelocity of about 17.0, based upon fresh feed, and a pressure of about860 p.s.i.g. The catalytic composite disposed in reactor 18, as well asin reactors 12 and 17, is a composite of alumina, 0.5% by weight oflithium and about 0.4% by weight of palladium. The effluent from reactor18 is withdrawn by way of line 19, passes through open valve 22 and line21, and, by way of line 16, is introduced into reactor 17 at a pressureof about 830 p.s.i.g.

Reactor 17 functions at a maximum `catalyst bed temperature of about 360F., representing an increasing temperature gradient of about F., and ata liquid hourly space velocity of about 1.8. With respect to the unitbeing illustrated, reactor 17 contains approximately 305 cubic feet ofcatalyst while reactors 18 and 12 each contain amout 34 cubic feet ofcatalyst. Thus, with respect to reactors 17 and 18, containing a totalof 339 cubic feet, reactor 18 contains about 10.0% of the total catalystin the two reactors. The effluent from reactor 17 is withdrawn by way ofline 23, and introduced thereby into heater 24, in admixture withmake-up hydrogen from line 26, wherein the temperature is increased to alevel of about 525 F. The heated mixture passes through line 25, and isintroduced into reactor 27 at a pressure of about 790 p.s.i.g. Reactor27 contains a catalytic cornposite of alumina, 5.0% by weight of nickeland 10.0% by weight molybdenum and in amount such that the liquid hourlyspace velocity therethrough, based upon fresh feed, is about 2.0.

The vincreasing temperature gradient `is maintained at a level of about75 F., resulting in a maximum catalyst bed temperature of 600 F. Thereactor efiiuent is withdrawn by way of line 28, and, following its useas a heatexchange medium to decrease its temperature to about 300 F., isintroduced into hot separator 29 at a pressure of about 770 p.s.i.g.

A principally liquid phase is withdrawn from hot separator 29 by way ofline 34, and about 3,721 bbl/day is diverted through line 2 to combinewith the fresh feed charge stock, in line 1. Where desired, a portion ofthe hot separator liquid may be recycled through line 34 to be combinedwith the regenerating hydrogen stream in line 33. A portion of the hotseparator liquid may be recycled to the third catalytic zone.

The principally gaseous phase from hot separator 29 is withdrawn by wayof line 30, and, following its use as a heat-exchange medium and furthercooling, is introduced into cold separator 31 at a temperature of about100 F. and a pressure of about 750 p.s.i.g. The hydrogen-rich recyclestream, for use in regenerating the catalyst in reactor 12, is withdrawnby way of line 33, and introduced thereby into heater 35. The normallyliquid hydrocarbons stream is withdrawn from cold separator 31 throughline 32, and is admixed therein with the excess liquid phase from hotseparator 29 in line 42, the mixture constituting the normally liquidproduct of the process.

Prior to being sent to an aromatic recovery system, the product in line32 is introduced into a hydrogen sulfide stripping column, the toptemperature of which is about 250 F., the top pressure Ibeing about 190p.s.i.g. and the bottoms temperature being about 430 F. At theseconditions, the bottom stream is recovered substantially free frombutanes and lighter hydrocarbons, and can be transported directly to thearomatic recovery system. With respect to as-produced aromaticconcentrate, the sulfur concentration is effectively nil being onlyabout 0.1 p.p.m. by weight. The diene value is nil, the bromine withre-running and depentanizing, the catalytic composite in reactors 17 and27 can be expected to attain a catalyst life of 140 and 150 `bbL/lb.,respectively; through the use of the present invention, the expectedultimate catalyst life is increased to about 280 and 300 bbL/lb.,respectively, and no re-running or depentanizing is required. Theerected cost of the reaction system, including heaters, where re-runningand depentanizing mustl be provided, is about $500,000.00. The erectedcost of a unit, incorporating the present invention, including heaters,is about $420,000.00, or about $30,000.00 per 1,000 bbl. of chargestock.

number indicative of the quantity of mono-olefins, is less than about0.3 and about 98.0% recovery of the aromatics, by weight, is attained.

Referring once again to the guard chambers, reactors 12 and 18, duringthe initial stage of regenerating catalyst in the former, valve 41 inline 40 will be open in order 5 to recover residual light liquidmaterial which can be introduced into reactor 27 by way of line 25. Whenthe gums and other oleins polymer appears in the eiuent fromknock-outpot 15, valve 41 is closed, valve 39 is open, and the polymerizationproducts are withdrawn from the process. When the catalytic compositedisposed in reactor 12 is substantially free from olen polymer productsand reactor 12 has been cooled to the process temperature, the followingvalves are opened: valve 4, line 3, valve 11, line 36, valve 8 in line10, valve 20 in line 19, and valve 17 in line 16; and, the followingvalves are closed in order to regenerate the catalyst in reactor 18:valve 7 in line 6, valve 9 in line 10, valve 5 in line 3, valve 14 inline 13, and valve 22 in line 21. The procedure is, of course, reversedafter the catalytic composite disposed within reactor 18 is regeneratedto ythe extent of being substantially free from olen polymer products.

In the absence of the swing-bed guard chambers but 'I'he foregoingspeciiication, and example integrated into the description of thedrawing, clearly illustrates the method of effecting the presentinvention and the benelits to be afforded through the utilizationthereof.

I claim as my invention: 1. A continuous process for hydrorening asulfurous,

(a) reacting said charge stock and a hydrogen-rich gaseous phase ashereinafter identified in a first catalytic reaction zone with acatalyst composite which is non-sensitive to sulfur and which is presenttherein in an amount of about 5.0% to about 25.0% by weight of the totalcatalyst in the rst and second reaction zones, in said first catalyticreaction zone at least partial hydrogenation of di-olelinic hydrocarbonsis eiected and olein polymer formation occurs, at a maximum catalyst bedtemperature in the range of about 200 F. to about 500 F.;

(b) further reacting the resulting first zone euent, at substantiallythe same temperature in the range of about 200 F. to about 500 F. in asecond catalytic reaction zone containing said catalyst compositewherein at least partial hydrogenation of di-oleiinic compounds issubstantially completed;

(c) increasing the temperature of the resulting second reaction zoneeffluent to a level above about 500 F., and reacting the thus-heatedsecond zone eiuent in a third catalytic reaction zone containing ahydrodesulfurization catalytic composite wherein monoolen hydrogenationand desulfurization are eiected at a maximum catalyst bed temperaturebelow about 1,000 F.;

(d) separating the resulting third reaction zone eiuent, 75

at a temperature in the range of about 200 F. to about 500 F., in afirst separation zone, to provide a rst vaporous phase and a firstliquid phase;

(e) further separating said first vaporous phase, at a substantiallylower temperature, in a second separation zone, to provide a secondliquid phase and a hydrogen-rich second vaporous phase;

(f) heating said hydrogen-rich second vaporous phase to a temperature inthe range of 500 F. to about 700 F.;

(g) passing the thus-heated second vaporous phase through a regeneratingguard chamber as a fourth catalytic reaction zone containing an olefinpolymerdeactivated catalytic composite;

(h) separating the resulting fourth reaction zone efuent, in a thirdseparation zone, to provide a polymer concentrate and a hydrogen-richthird vaporous phase; and,

(i) recycling said third vaporous phase to combine with said chargestock as said hydrogen-rich gaseous phase in step (a).

2. The process of claim 1 further characterized in that said rst andfourth catalytic reaction zones are interchanged when the deactivatedcatalytic composite in said irst reaction zone is substantially freefrom olefin polymers, so that said charge stock and regeneratinghydrogen gaseous phase are introduced into said fourth reaction zone,and said heated second vaporous phase is introduced into said firstcatalytic reaction zone to remove olefin polymers from the catal'y'stdisposed therein.

3. 'I'he process of claim 2 further characterized in that said first andfourth reaction zones are interchanged when the catalyst in said firstreaction zone is substantially free from olelin polymers.

4. The process of claim 1 further characterized in that at least aportion of said rst liquid phase is recycled to combine with said chargestock.

5. The process of claim 1 further characterized in that at least aportion of said tirst liquid phase is recycled to combine with saidsecond vaporous phase.

6. The process of claim 1 further characterized in that at least aportion of said rst liquid phase is recycled to said third catalyticreaction zone.

7. The process of claim 1 further characterized in that the catalystdisposed in said iirst, second and fourth reaction zones is a compositeof a Group Vllll noble metal component, an alumina-containing non-acidiccarrier material and an alkali metal component.

8. The process of claim 1 further characterized in that the catalyst bedtemperature in said third reaction zone is in the range of 500 F. toabout 800 F. and the catalyst disposed therein contains an iron-groupmetal component.

l9. The process of claim 1 further characterized in that the maximumcatalyst bed temperature in said third reaction zone is in the range of800 F. to about l,000 F. and the catalyst disposed therein contains aGroup VIII noble metal component.

References Cited UNITED STATES PATENTS 3,457,163 7/ 1969 Parker 208-2112,833,698 5/ 1958 Patton et al. 208-211 3,239,449 3/1966 Graven et al.208-210 2,833,697 5/1958 Oettinger 208-211 3,167,498 1/ 1965 Leverkusenet al. 20S-210 3,492,220 1/ 1970 Lempert et al. 20S-210 DELBERT E.GANTZ, Primary Examiner G. I. CRASANAKIS, Assistant Examiner U.S. C1.XJR. 208--97

